Azeotropic distillation procedure



April 5, 1966 w. D. scr-IAEFFERv ET AL 3,244,759

ZEOTROPIC DISTILLATION PROCEDURE Filed OCT.. 5l. 1961 05754 mafia United States Patent O 3,244,759 .A AZEOTROPIC DlSTILL'ATION' PROCEDURE William D'. Schaeffer, Pomona, and Vernon E. Stiles,

Fullerton, Calif., assignors to Union Gil-Company of California, Los Angeles, Calif., a corporation of California Filed Oct. 31, 1961, Ser. No. 149,100 12 Claims. (Cl. 260-672) This invention relates to the manufacture of naphthalene and Z-methylnaphthalene, and .in particular concerns an improved process-for treating certainl hydrocarbon mixtures of petroleum origin to recover therefrom, as process products, both naphthalene and Z-methylnaphthalene'of a high degree of purity. One point of departure of our process from conventional, ,or other, dealkylation procedures is the use of azeotropic distillation as a means of recovering'Z-methylnaphthalene in high purity from certain alkylnaphthalene heart-cut fractions of dealkylation recycle streams. The origin, identity and nature of the aforesaid dealkylation yrecycle streams will be apparent from the complete description of our invention to follow.

It is well known that certain hydrocarbon mixtures of petroleum origin, as exemplified by heavy reformate fractions, contain appreciable amounts of methylnaphthalenes, and that such mixtures can be catalytically or thermally treated to effect dealkylation of said methylnaphthalenes and'thus obtain naphthalene-as a product. In carrying out dealkylation processes of this type, it has been discovered that l-rnethylnaphthalene dealkylates from 2 to 5 times faster than-Z-methylnaphthalene withfthe result that there is a buildup of the latter isomer inprocess recycle streams. Becausev Z-methylnaphthalene is a potentially valuable byproduct, and its presence in dealkylation process recycle streams is detrimental (because of its relative slowness to dealkylate and its consequent tendency tov accumulate in the stream), a means ofisolating it as a high grade (e.g., 95 percent pure) product would-obviously serve a double purpose. However, in the past the primary interest has been-in producing max-imurnamounts of naphthalene from the feedstock and until recently the possibility of removing Z-methylnaphthalene to accomplish the aforesaid twofold purpose was neither considered nor exploited, insofar as we are aware.

As a result of the present invention, the dealkylation ofl 1- and Z-methylnaphthalene and dimethylnaphthalene containing mixtures may now be conducted in such fashion as to effect conversion of a large proportion of the l-methylnaphthalene and the fdimethylnaphthalene to naphthalene (the latter undoubtedly pass through the monomethylnaphthalene stage on their way to naphthalene) while simultaneously providing highly v efficient means for removing Z-methylnaphthalene from the sys- Item as a high-grade by-product.

It is thus a principal object of our invention to provide a practical and eilicient process for the dealkylation of mixtures of the above-indicated type whereby the abovestated objectives can be realized.

Other objects, features and advantages of our invention will become apparent to thoseV skilled in the art as the description of the invention proceeds.

The expanding use of naphthalene for the production of dicarboxylic acids useful in manufacturing synthetic resins and fibers has created considerable interest in the manufacture of naphthalene from petroleum hydrocarbons. It is well known that certain hydrocarbon fractions obtained in various petroleum refining operations, for example, cracking and reforming, contain considerable quantities of alkylnaphthalenes which can be thermally or catalytically dealkylated to form free naphthalene. suchrprocesses are conventionally carried out, a feedstock Asy consisting of an alkylnaphthalene concentrate, boiling above about 430 F., from a reformate or cycle oil Yfraction which comprises alkylnaphthalenes and such non'- naphthalenic materials as alkylbenzenes, alkyltetralins, alkylindanes, etc., is subjected to dealkylating conditions to obtain an efluentA product which, after separation of normally gaseous materials such as hydrogen and low molecular weight hydrocarbons, is fractionally distilled to obtain: (l) a light gasoline fraction, (2) a naphthalene fraction, and (3) a heavy fraction comprising unreacted alkyinaphthalenes which, after the removal of heavy ends and polymers, is recycled to the reactionv zone. According tol one mode of operation,th'e` naphthalene fraction is taken over a relatively wide range, foriexample', from about 400 F. to about 435 F andthe naphthalene is recovered therefrom in substantially pure formA by'low temperature crystallization or vazeotropic distillation.

Where the dealkylation feedstock contains a significant quantity of 2-methylnaphthalene, as such feedstocks normally do, and-the above-describedprocedureyis followed there is a buildup of recycling 2-methylnaphthalene in the system for previously indicated reasons. However, when the dealkylation method is carried out according' to the improved procedure of this invention the buildup of 2- methylnaphthalene vis prevented andl in addition the 2- methylnaphthalene is recovered in relatively pure form' as a potentially valuable product. Brieiiy, in carryingout the dealkylation method according to our improved procedure, the dealkylation reactor eiuent (consisting of liquid hydrocarbons, obtained as indicated above and described in greater detail below) is distilled to obtain an efliuent naphthalene fraction, a' monom'ethylnaplithal'enel fraction, and a'higher boiling fraction comprising polymethylnaphthalenes and' other alkylnaphthalenes. The process feedstock is similarly'distilled to obtain a monomethylnaphthalenefraction.- Usually, thev feedstock will be of fairly wide boiling range; comprisingnaphthalene and/ or alkylnaphthalenes of higher boiling point than the monomethylnaphthalenes,and in'such event, the feedstock is distilled to isolateithese" components as fractions of relatively narrow boiling range;

In accord-ance with-'a'preferred mode of operation, a relatively wide boiling-range process feedstock is admixed with the aforesaid reactor effluent andthe combined mixture'is distilled tov obtain the aforesaid fractions. The ynaphthalene fraction (whether it be obtained solely from the reactor efluent or from both thereactor effluent and the processv feedstock, distilled together orseparately) is treated byv fractionation means such as azeotropic distillation, or crystallization and centrifuging to obtain essentially pure naphthalene and a filtrate or equivalent-by-product stream which is useful as a heavy gasoline blending stock. The monomethylnaphthalene fraction (whether it be obtained by distilling the reactor effluent and the feedstock together or separately) is subjected to azeotropic distillation in the presence of an azeotrope former, such as monoethanolamine or the like, to obtain as an overhead an azeotropic distillate comprised of the azeotrope former and non-naphthalenic hydrocarbons from the monomethylnaplithalene fraction and as a bottoms product substantially pure 2- met-hylnaphthale'ne. product is separated as a relatively pure by-product of the overall process. The azeotropic distillate overhead is' condensed and then separated into two liquid phases-Y one comprising substantially pure azeotfrope former and the other compsingrnon-naphthalenic hydrocarbons removed fromv the monomethylnaphthalene fraction. The azeotrope former fraction from the phase separation operation is recycled to t-he'azeotropic distillation column' andthe non-naphthalenic hydrocarbon fraction can be ad- The Z-m'ethylnaphthalene bottoms mixed with the higher boiling alkylnaphthalene fraction (whether it be obtained solely from the reactor effluent or from both the reactor efliuent and the process feedstock, distilled together or separately) as feed to the dealkylation reactor or otherwise disposed of.

Within the dealkylation reactor, the hydrocarbons are contacted with an alkylated catalyst comprising cobalt and molybdenum oxides, the contacting being effected in the presence of added water and hydrogen at a temperature of about 9501250 F. and at -a pressure of about 600- 1500 p.s.i.g. The effluent stream from the reactor is cooled and condensed to separate hydrogen and normally gaseous hydrocarbons, the water is decanted from the liquid phase, and the hydrocarbon phase is distilled as previously described. We have found that by following the foregoing procedure it is possible -to convert to naphthalene as much as 85% or more of the 1methylnaph thalene and alkylnaphthalenes boiling above the monomethylnaphthalenes, while converting as little as 35% or less of the 2-methylnaphthalene, and at the same time recovering as much as 60% or more of the unconverted 2-methylnaphthalene in substantially pure form.

Reference is now made to the drawing, which forms a part of this application and takes the form of a diagrammatic flowsheet illustrating a preferred mode of carrying out t-he process of our invention. In the interests of simplication, such conventional process equipment as pumps, heat exchangers, reboilers, condensers, reilux lines, phase separators, valves, instrumentation, and so forth, have, in many places, been omitted from the drawing.

Turning now to the drawing, the description thereof is set forth in terms of a specific operating example employing a process feedstock of the following characteristics:

Source Platinum catalyzed reformate fraction. Gravity 19 API. Boiling range 379630 F. Alkanes 2.4 wt. percent. Naphthenes 0.4 wt. percent. Alkylbenzenes 27.1 wt. percent. Tetralins-t-indanes 22.0 Wt. percent. Other non-naphthalenics 3.2 Wt. percent. Naphthalene 9.8 wt. percent. l-methylnapht-halene 5.8 wt. percent. Z-methylnaphthalene 12.2 wt. percent. Dimethylnaphthaleues 12.7 wt. percent. Higher naphthalenics 4.3 wt. percent.

As will be noted, the major items of processing equipment include light gasoline column 10, naphthalene column 12, methylnapht'halene column 14, methylnaphthalene overhead condenser 100, reactor feed column 16, reactor feed lheater 18, catalytic reactor 20, reactor effluent cooler 22, high pressure reactor etliuent separator 24, low pressure separator 23, naphthalene recovery system 26, and a Z-methylnaphthalene recovery system comprising azeotropic distillation column 28, azeotropic distillate condenser 29, azeotropic distillate phase separator 27, and azeotropic distillation column 25 in whic-h the non-naphthalenic-rich fraction from phase separator 27 is azeotropically distilled to remove minor amounts of azeotrope former entrapped therein.

The feedstock enters the process system via line 30 and is joined therein by a reactor eiuent stream flowing in line 32 and comprising liquid hydrocarbon components of the reactor efduent. From line 30 the mixture of feedstock and reactor etiluent is introduced into gasoline column 10, which suitably takes the form of a 25-tray bubble-cap column operated to produce an overhead fraction boiling between about 225 F. and about 405 F. This overhead fraction, comprising C-C9 hydrocarbons formed in the dealkylation reaction as well as those present in the feedstock, is passed via line 33 to storage for use as light gasoline or otherwise. The bottoms fraction from column 10 is passed via transfer line 34 to naphthalene column 12, which is suitably a 30-tray bubble-cap column operated to producel an overhead naphthalenecontaining stream boiling between about 405 F. and about 425 F. The bottoms fraction from column 12 is passed via transfer line 36 to methylnaphthalene column 14, which is suitably a 30-tray bubble-cap column operated to produce an overhead methylnaphthalene-containing stream boiling between about 425 F. and about 460 F. The bottoms fraction from column 14 is passed via transfer line 38 to reactor feed column 16, which suitably takes the form of a 20-tray bubble-cap column operated to produce an overhead fraction boiling between about 460 F. and about 520 F., and a bottoms fraction which -is passed via line 40 to storage as fuel oil or the like. The overhead fractions from columns 12 and 14 are passed via lines 42 and 44, respectively, to naphthalene recovery system 26 and to azeotropic distillation column 28 (through condenser 100), respectively. The operation of these recovery systems is described in detail below.

The overhead fraction from reactor feed column 16 is passed via line 46 to reactor feed manifold 48. The co1- umn 16 overhead is passed by manifold 48 to a reactor feed heater 18, but prior to entering that heater it is joined by the bottoms fraction from distillation column 25 through line 49 as shown on the drawing. The bottoms fraction from column 25 comprises the non-naphthalenic hydrocarbons removed from the column 14 overhead in the manner hereinafter more fully described. The column 16 overhead is also joined, prior to passing into reactor feed heater 18, by a mixture of water vapor and a hydrogen-containing recycle gas stream flowing in line S0. Makeup hydrogen and water are introduced into line 50 from lines 76 and 51, respectively. The total reactor charge is passed to reactor feed heater 18 wherein it is heated to an incipient reaction temperature, and the heated reactor feed is passed via line 1 to catalytic reactor 20 at such rate that the space velocity within the reactor is about 0.4 volume of hydrocarbon per hour per volume of catalyst. About 8100 s.c.f. of hydrogen and about 0.085 barrel of water are provided per barrel of hydrocarbon in the reactor feed mixture.

Reactor 20 contains a fixed bed of a dealkylation catalyst consisting of a silica-stabilized alumina base having supported thereon about 3.0 weight percent of cobalt oxide, about 9.0 weight percent of molybdenum oxide, and about 2.2 weight percent of sodium oxide. The reaction zone is maintained at a temperature of about 1100 F. and a pressure of about 1000 p.s.i.g., and a substantially uniform temperature profile within the catalyst bed is maintained by introducing a quench gas consisting of a part of the hydrogen-containing recycle gas owing in line 50 into the reaction zone at a plurality of points along the length thereof via lines 2, 3, 4 and 5. About 1430 s.c.f. of the quench gas are employed per barrel of hydrocarbon in the reactor feed. The reactor eluent passes from reactor 20 through line 54 and into effluent cooler 22 wherein it is cooled to about 160 F., and thence into high pressure separator 24 which is operat-ed at a pressure only slightly below that of reactor 20. Within separator 24, the cooled etliuent separates into a gaseous-I phase, a liquid hydrocarbon phase, and an aqueous phase. The latter is withdrawn via line 56 and is either discarded' or recycled back to the reactor feed manifold. The gaseous phase, which comprises hydrogen and a relatively small amount of normally gaseous hydrocarbons, is withdrawn from separator 24 via line 50 for use as the reactor quench gas and for recycle to the reactor feed manifold.. The liquid hydrocarbon phase in separator 24 is withdrawn via line S0 and is passed through pressure relief valve S2 into low pressure separator 23 which is operated at or near atmospheric pressure. Within separator 23, the dissolved normally gaseous hydrocarbons flashv olf to form a gaseous phase which is withdrawn and passed via line 58 to storage for use as fuel gas or the like. The liquid hydrocarbon phase which forms in separator 23y is introduced into line 32 for return to .column 10 and ultimate recovery of gasoline, naphthalene, and 2methylnaphthalene therefrom. The small amount of water which separates in separator 23 is passed to waste.

Returning now to the product recovery systems, the naphthalene concentrate taken overhead from naphthalene column 12 is passed via line 42 to naphthalene recovery system 26, which is preferably a conventional crystallizercentrifuge combination. The crystallizer is suitably a 'scraped surface type in which the naphthalene concentrate is chilled to a temperature of Yabout F. over a period of time between about 30 and about 60 minutes to permit the crystallization of solid naphthalene. The latter is scraped from the cooling surfaces of the crystallizer, and is transferred to the centrifuge wherein it is separated from the mother liquor. Suitably, the crystallized naphf thalene is washed in the centrifuge with a light hydrocarbon wash liquid such as pentane.` The naphthalene product is dried and passed to storage via line 66, and the mother liquor (after removal of the wash liquid, for example by distillation) is passe-d via line 62 to storage for use as a heavy gasoline blending stock or the like.

An azeotropic distillation system can be employed as an alternate to, or in conjunction with, the above-described crystallizer-centrifuge method of naphthalene recovery if desired. In the former event, the column 12 overhead is condensed and subjected to distillation in they presence of a suitable azeotrope former, such as monoethanolamine, whereby an azeotropic distillate of non-naphthalenic hydrocarbons and azeotrope former is removed, leaving a bottoms which can be recovered as the naphthalene product of the overall process. The non-naphthalenic hydrocarbons are separated from the azeotropic distillate and stored for use as a heavy gasoline blending stock or otherwise disposed of. i

The foregoing azeotropy disclosure is oversimplied but believed adequate for present purposes for several reasons. For one thing, the azeotropic distillation of the naphthalene concentrate from column 12 is similar procedurewise to that of the methylnaphthalene concentrate from column 14 and the latter procedure is described in detail infra. The similarity extends even to the identity of the azeotrope formers, the same materials being suitable as azeotrope formers in either procedure. Furthermore, cross reference will be hereinafter made to copending U.S. patent applications describing in detail the azeotropic distillation of like boiling non-naphthalenic hydrocarbons from naphthalene containing mixtures by the subject procedure.

Returning now to the drawing, the overhead from methylnaphthalene distillation column `14 is passed via line 44 to condenser 100, wherein it is liquied, and from thence lto azeotropic distillation column 23. Column 28 may be a conventional fractional distillation column containing, for example, -80 plates and may be operated at overhead reflux ratios of, for example, 1:1 to 20:1. Azeotroping is continued until essentially -all of the nonnaphthalenic hydrocarbons 'boiling within about 10 F. and preferably 20 F., of Z-met-hylnaphthalene, are distilled overhead. The overhead temperatures during azeotropic distillation will preferably range between about 330 and 338 F., the fazeotrope former itself preferably boiling at -around 340 F. The azeotropic overhead is taken off through line 88 and condensed in condenserE 29, and the resulting condensate is transferred via line 90 to a liquid phase separator 27, wherein two phases, an azetrope former phase and non-naphthalenic hydro- 'carbon phase, stratify. The azeotrope former phase which straties in separator 27 will ordinarily contain about 0.5 to about 1.0 percent of dissolved hydrocarbons, and this solution is continuously recycled via line 6 6 back to a midpoint in distillation column 2,8.. To maintain the desired reflux ratio, a portion of the condensate in line ,may be Y,diverted through line 7 and admitted to the top of Vcolumn 28. Makeup azeotrope former can, of course, be added ,to the system, by means not shown on the drawing, if desired.

The non-naphthalenic hydrocarbon-rich phase in separator 27 ordinarily will contain about 0.1 to about 0.5 percent by weight of dissolved azeotrope former. This hydrocarbon-rich 4phase is then transferred via line 8 to fractionating column 25, from which substantially all of the azeotrope former is distilled overhead via line 9 as an azeotrope with a portion Aof the non-naphthalenic hydrocarbons, and condensed into separator 27. Nonnaphthalenic hydrocarbons are withdrawn as bottoms from column 25 via line 49. The bottoms fraction from column 25 is passed to manifold 48 for use as previously described. The Z-,methylnaphthalene bottoms fraction from column i28 i s withdrawn through line 92 and passed to storage or other disposition not shown on the drawing.

The following tabulation sets Yforth operating conditions and other details -of a typical operation of the process carried out as ldescribed above, using a conventional crystallizer-centrifuge combination as the naphthalene recovery system, with the aforementioned process feedstock. All boiling ranges are ASTM. For purposes of illustration, the present tabu-lation presumes the use of monoethanolamine as the azeotrope former in the 2- methylnaphthalene recovery system.

Hydrocarbon feed rates:

Process feedstock (line 30) lbs./hr. 974 Reactor effluent (line 32) lbs./hr. 376

Gasoline column operation:

Overhead fraction ra-te (line 33) lbs./hr Overhead fraction boiling range, F. 200-405 Naphthalene column operation:

Overhead fraction rate (line 42) lbs./hr 684 Overhead fraction boiling range, F. 405-427 Overhead fraction, wt. percent naphthalene 26.5

Methylnaphthalene ,column operation:

Overhead fraction rate (line 44) lbs./hr 286 Overhead fraction boiling range, F. 427-454 Overhead fraction, wt. percent 2methylnaphthalene 66 Overheadk fraction, wt. percent l-methylnaphthalene 22 Reactor feed column operation:

Overhead fraction rate (line 46) lbs/hr-- 192 Overhead fraction boiling range, F. 454-507 Bottoms `fraction rate (line 40) lbs./hr 68 Overhead fraction Iboiling range, F. 507-754 Reactor charge:

Reactor feed column overhead rate (line 46) lbs./hr 192 Non-naphthalenic bottoms fraction from column 25 (line 49) lbs./hr 142 Water, percent by vol. of hydrocarbons 8.5 Recycle gas (90% hydrogen), s.c.f./=bbl.

hydrocarbon 11,000

Reactor conditions:

Reaction temperature, average 1099 F. Reaction pressure, p.s.i.g 1000. Liquid hourly space velocity 0.4.

Catalyst:

9.0 wt. percent molybdenum oxide +30 Wt. percent cobalt oXide-l-2-2 wt. percent sodium oxide supported on gamma alumina. stabilized with wt. percent silica, 1s-inch pellets.

Quench gas (90% hydrogen)rate, s.c../ bbl.

hydrocarbon 1430.

Naphthalene recovery system:

Feed rate (line 42) lbs./hr 684 Chiller temperature, F. -10 Residence time minutes 45 Wash liquid Pentane Naphthalene produced lbs./hr 155 Heavy gasoline produced lbs./hr 529 2-methylnaphthalene recovery system:

Feed rate, column 14 overhead lbs./hr. 286 Feed rate, monoethanolamine recycle lbs./hr. 165

2-methylnaphthalene produced lbs./hr 144 Column 25 bottoms fraction rate (line 49) lbs./hr 142 It will be apparent to those skilled in the art that yields and purities of our 2-methylnaphthalene product will vary directly with the degree of l-rnethylnaphthalene selectivity in the dealkylation zone, the higher the degree of such selectivity, the greater the Z-methylnaphthalene yields and purities. In this connection, the above-identied catalyst has been found to exhibit a high degree of selectivity for l-methylnaphthalene dealkylation, evidence of which, in form of experimental data, is offered below. These data were obtained by subjecting each of three different feedstocks to a catalytic dealkylation in accordance with the process of the invention. Feed No. 1 was a iplatinum catalyzed reformate bottoms fraction boiling over the range 400 F.-550 F. Feed No. 2 likewise had a boiling range of 400 F.-550 F., and was composed of 57 volume percent of Feed No. 1 and 43 volume percent of a product obt-ained by subjecting Feed No. 1 to catalytic dealkylation and removing the naphthalene and lighter materials from the hydrocarbon reactor eflluent. Feed No. 3 was a product obtained by subjecting Feed No. 1 to azeotropic distillation to concentrate the naphthalenic components. Reaction condition for all three runs were as follows:

Temperature, F. 1000.

Pressure, p.s.i.g 880.

Liquid hourly space velocity 0.99 vol. hydrocarbon/- lin/vol. catalyst.

Water in reactor charge 24 vol. percent of hydrocarbon. Hydrogen in reactor charge, 6252 s.c.f./bbl. hydrocaraV. bon. Length of run hours.

Catalyst:

3.0V wt. percent COO supported on gamma-alumina, 9.0 wt. percent M003 stabilized with 5 wt. percent, 2.2 .wt. percent Nag() silica, 1s-inch pellets.

Feed Feed 1 Feed 2 Feed 3 Analysis, percent by Wt.:

5 urates 5. 5 1. 2 0.8 Alkylbenzencs 17. 5 13. 5 8. 4 Tetralins-lindanes 20. 9 19. 8 15. 5 Naphthalene 5.1 7.1 4. 9 1-rnethylnaphthalene 10. 4 8. 9 14. 6 Z-methylnaphthalene 21. 5 31. 9 29. 5 Other naphthalenics 17. 2 15. 4 23. 3 Higher fused rings 1. 9 2. 2 2. 9

PRODUCT Analysis, percent by wt.:

(J1-400 F. fraction 34. 7 26. 1 21.1 400 Pfl-fraction:

Saturates U. 8 0. 9 0. 3 Alkylbenzenes 5. 2 5. 0 4. 0 Tetralins+iudanes 10. 5 11. 0 9. 5 Naphthalene 22. 2 24. 2 26. 2 lmethylnaphthalene 2. 7 2. 4 4. 8 Z-methylnaphthalene 19. 8 26. 2 27. 2 Other naphthalenics. 4. 1 4. 2 6. 9 Higher fused rings 0. 0 0. 0 0. 0 1-methylnaphtl1alene consumed, percent 75. 5 73. 0 67. 1 2-methy1naphthalene consumed, percent- 7. 9 17.9 7. 8

Ratio: 1methylnaphthalene consumed/2- methylnaphthalene consumed 9. 6 4. 1 8.6

In contrast to the foregoing, the catalytic dealkylation process described in U.S. 2,943,120, a suitable standard for present comparative purposes, the ratio of 1-methylnaphthalene consumed to 2-methylnaphthalene consumed varies trorn only about 1.42 to only about 2.92.

by changing the quantity of Water present in the reactor charge. Other reaction conditions were as follows:

Temperature, F 1000.

Pressure, p.s.i.g 750-880 Welght hourly space velocity... 0.905t 1gms. hydrocarbon/hr./grus.

ea a yst.

Hydrogen in reactor charge, av 694-4 sci/bbl., hydrocarbon.

Length of run 10 hours.

Catalyst:

3.0 Wt. percent C00 supported on gamma-alumina. 9.0 wt. percent M003 stabilized with 5 wt. percent. 2.2 wt. percent N a20 silica, %inch pellets.

Run No 1 I 2 s l 4 Partial pressure H2O at reactor inlet,

p.s.l.g 0 5l 82 163 l-methylnaphthalene consumed, percent. 88. 7 76. 4 74. 5 69. 8 2-methylnaphthalene consumed, pel'eent 34.1 19. 0 18. 1 9. 5 5() Ratio: l-rnethylnaphthalene consumed] 2.61 4. 01 4.10 7. 34

2-methylnaphthalene consumed It will be apparent to those skilled in the art that there are a great n-umber of ways of carrying out the various phases of the improved process of this invention. For example, where it is desired to take advantage of the selective effect of Water in the dealkylation reaction zone, this can be accomplished by introducing water into said reaction zone at one or more points along the length thereof. This serves both to increase the dealkylation selectivity and as a quenching medium to maintain satisfactory control of the reaction temperature. For a more detailed description of this particular method of carrying out the dealkylation reaction, see our copending U.S. paent application Serial No. 117,553, led June 16, 1961, now abandoned.

Considering now the improved dealkylation process of our invention in somewhat greater detail, the process feedstock may be any hydrocarbon mixture comprising a substantial amount, e.g., 10 weight percent or greater, of mono-rnethylnaphthalene isomers. Usually, however, it will be one of petroleum origin obtained as a heavy fraction from cracking or reforming operations, and will contain 2-50 weight percent of 1-methylnaphthalene, 2-60 weight percent of Z-metnylnaphthalene, and 2-50 9 weight percent of higher alkyln-aphthalenes and polymethylnaphthalenes. Catalytically or thermally cracked cycle oils and heavy reformate fractions are most suitable, with platinum-catalyzed reformate fractions boiling between about 375 F. and about 650 F. being particularly preferred. A second preferred feedstock consistsv `of a thermally or catalytically cracked cycle oil extract fraction obtained, for example, by subjecting a TCC or FCC cycle oil (boiling range=375650 F.) to hydrore'tining (i.e., hyd-rodesulfurization and/or hydrodenitro'- genation) in the presence of a cobalt molybdate catalyst (Unining), and subjecting the hydrore-ned product to solvent extraction (e.g., with sulfur dioxide or a glycol) to obtain an extract product rich in naphthalene an-d naphthalen'e precursors. Alternatively, the cycle oil can first be solvent extracted, and the aromatic's-rich extract then hydr-orefned to obtain a suitable feedstock. With particularly clean cycle oils (i.e., low sulfur and/or low nitrogen) the hydrorening step can be omitted.

The hydrocarbon portion of the reactor change cornprises mono-methylnaphthalenes, higher alkylnaphthalene's and polymethylnaphthalenes contained in the process feedstock and in the reactor efliuent, and mono-methylnaph'thalenes contained in the non-naphthalenic bottoms product from the Z-methylnaphthalene recovery system. When employing the preferred process feedstock this portion of the reactor charge typically boils within the range of about 440 F. to about 525 F., and contains Iabout 8-20 weight percent of Z-methylnaphthalene and about -30 weight percent of l-methylnaphthalene.

In order to prevent uncontrolled cracking it is necessary that water Vapor be present in the reaction zone in an amount corresponding to at least 0.04 barrel of liquid water per barrel of hydrocarbon. The water vapor can be introduced as such directly into the reaction zone at the upstream end of the catalyst bed, but more preferably it is admixed with the hydrocarbon feed prior to introduction of the same into the reaction zone. In a particularly preferred mode of operation liquid water is mixed with liquid hydrocarbon and the mixture is vaporized and preheated before being passed into the reaction zone. When following such procedure, between about 0.04 and about 0.15 barrel of liquid water is employed for each barrel of liquid hydrocarbon introduced into the reaction zone. It is also preferred to reduce the amount of water as the run proceeds and the activity 'of the catalyst decreases. Thus, it is preferred that at the beginning of a run using fresh or freshly regenerated catalyst, the reactor charge contain 0.10-0.15 barrel of water per barrel of hydrocarbon and that the proportion of Awater be thereafter decreased so that at the end of the effective life of the catalyst about 0.04 to 0.10 barrel of water is employed per barrel of hydrocarbon. In a typical operation, the volume ratio of water to hydrocarbon in the reactor charge is about 0.12/ 1 at the start of a run with new or freshly regenerated catalyst, and after about 5 days operation this ratio is reduced to about 0.08/1 yand is maintained there for the remainder of the effective active life of the catalyst.

In addition to hydrocarbons andwater vapor, hydrogen is introduced into the reaction zone as part of the reactor charge in an amount which may be varied between about 4000 and about 12,000 s.c.f./bbl. of hydrocarbon and which is preferably between about 6000 and about 10,000 s.c.f./bbl. Preferably the hydrogen is employed in the form of an enriched recycle gas containing at least'75, preferably 85-90, volume percent of hydrogen. The net consumption of hydrogen in the reaction zone amounts to about 400-800 s.c.f. per barrel of hydrocarbon introduced into the reaction zone.

The dealkylation catalyst comprises an activated alumina support upon which is distended a minor amount of molybdenum oxide (e.g., 5 tov 20 weight 4perCenUand a smaller amount of cobalt oxide (eig, 1 to 5 weight percent). Such mixture of cobalt and molybdenum oxides is frequently referred to as cobalt molybdate even though the two oxides may not be present in equimolecular proportions. It is preferred that the catalyst contain a stabilizing anio'nnt off silica (e.g., 3 to l5 weight percent), and a mnior amount (ejg. 0.l to 3 weight percent) of an alkali-metal or an alkaline earth metal. The latter metals are conveniently added to the uncalcined composition in theform of a hydroxide, carbonate or other heatdecomposable compound 'so that upon final calcination the alkalizing component is converted to an oxide, hydroxide, silicate, aluminate or the like. To prepare catalysts of this type, a hydrogel of the carrier and stabilizer is irst prepared, as by precipitation or coprecipitation from aqueous solution. The hydrogel is then dried and impregnated with soluble salts of cobalt land molybdenum, e.g., cobalt nitrate and ammonium molybdate, either ysimultaneously or alternately. The alkalizing cornponent, eg., sodium hydroxide, sodium carbonate, potassium carbonate, calcium hydroxide, etc., is then added in a final impregnation step, after which the composition is dried and 'calcined in the conventional manner. Alternatively, the alkalizing component may be added prior to yor simultaneously with one or both of the heavy metal components, eg., it may be lmxie'd with an ammonium molybd-ate solution and impregnated simultaneously with lthe molybdenum.

The reaction temperature is between about 950 F. and about 1250 F., with temperatures at the lower end of said range being employed with fresh catalyst of maximum activity and the higher temperatures being employed as the catalyst activity decreases. The reaction temperature is maintained by heating the 4reactor feed mixture to a temperature between about 950 F. and about l200 F. before introducing it into the reaction zone. Suitable reaction pressures ran'ge from about 600 to about 1500 r pounds per square inch gauge, with pressures of about 900 to about 1200 pounds per square inch gauge being preferred. The liquid hourly space velocity may vary from about 0.2 to about 5 volumes of hydrocarbon per hour per volume of catalyst, with 0.3 to 0.8 being pre'- ferred. f

Since the dealkylation reaction is highly exothermic, it is necessary to provide means for preventing the development of excessive temperatures within the reaction zone. This is readily accomplished by introducing a quenching medium into the reaction z'one at such points along the length thereof and in such amounts as to maintain the temperature rise along the length of the reaction zone at a value less than F. Either hydrogen or water vapor may be employed as the quenching medium. Hydrogen is preferred when it is desired to control the process to produce a maximum amount of naphthalene, whereas water vapor is preferred when it is desired to produce a maximum amount of Z-methylnaphthalene. In either case the amount required to secure the necessary limitation in temperature rise will depend upon the identity of the hydrocarbon portion of the reactor charge, the temperature of the quenching medium, the activity of the catalyst, the number and proximity of the points at which the quenching medium is introduced into the reaction zone, etc. Also, the amount of quenching medium required will be greatest at the start of a run when the catalyst is of highest activity. Typically, 'which hydrogen being employed as the quenching medium (preferably in the form of the aforementioned enr'ichedrecycle gas), about 1500-2000 s.c.f./bbl. of hydrocarbon introduced into the reaction zone are required at the beginning of a run, with this amount being reduced to about 1000-1500 s.c.f./bbl. by the time the catalyst is ready for regeneration. When water is employed as the quenching medium it is typically provided in an initial amount corresponding to about 0.05-0.10 barrel of water per barrel of hydrocarbon introduced into the reaction zone, with this amount being reduced to VVeration.

about 0.04--08 barrel of water per barrel of hydrocarbon at the end of the run.

Those skilled in the art will understand that the foregoing operative ranges of reaction conditions and reactant proportions should be properly correlated to obtain optimum results.

Mention has perviously been made of catalyst regen- As is the case with substantially all processes involving the catalytic conversion of hydrocarbons, the present process is characterized by a gradual reduction in catalyst activity so that eventually it becomes necessary to renew or regenerate the catalyst. in general, regeneration of the present catalysts follows conventional lines Well known to those skilled in the hydrocarbon conversion arts. The aforementioned preferred catalysts are especially well adapted to regeneration since their relative insensitivy to water allows water vapor to be included as a component of the regeneration gas for the purpose of absorbing the exotherrnic heat of regeneration.

Procedure-wise, the process of the invention may be carried out simply by preheating the reactor charge (comprising hydrocarbon, water and hydrogen) to an incipient reaction temperature, and passing the vaporized charge through a unitary fixed bed of catalyst while injecting the quench medium into the catalyst bed at a plurality of points more or less uniformly located along the length thereof. Sometimes, however, it is preferred to divide the entire catalyst charge into two or more, preferably four, separate beds and to introduce the quenched medium into the system between adjacent beds. For a more detailed description of this multi-bed type dealkylation operation, see our previously mentioned copending application of Serial No. 117,553.

It is normally preferred, particularly when the process feedstock contains an appreciable amount of light gasoline, to recover such gasoline along with the light gasoline contained in the reactor effluent by combining the two streams and recovering all of the gasoline in a single distillation column as shown rin the drawing. Similarly, when a process feedstock contains an appreciable quantity of naphthalene, the latter is preferably recovered simultaneously with the naphthalene contained in the reactor effluent by a similar combined distillation, also as shown on the drawing. Thus, the point at which the process feedstock in introduced into the system will frequently depend upon its composition. If desired, of course, the process feedstock may be treated as an individual stream to obtain a material suitable for direct introduction into the reaction zone. While it is preferred to treat both feedstock and reactor eflluent material in one distillation operation as described for the removal of light gasoline as well as one operation for the removal of naphthalene, it is of course within the scope of our invention and sometimes even preferred to separately treat these two streams to remove gasoline and naphthalene therefrom.

The reactor efliuent comprises unconverted components of the feedstock, naphthalene, light gasoline, normally gaseous hydrocarbons, water vapor and hydrogen. The normally liquid components are separated by condensation, with the condensed water being separated from the liquid hydrocarbons by settling and decantation (preferably at a somewhat elevated temperature, eg., l50-200 F., to assure good phase separation) and the normally gaseous components are preferably in part employed (after enrichment with hydrogen) to provide the hydrogen component of the reactor charge. The liquid hydrocarbons are then distilled, either as such or in admixture with the process feedstock as described above, to obtain a light gasoline fraction, a naphthalene concentrate fraction, a mono-'methylnaphthalene fraction (which is treated as herein described to recover Z-methylnaphthalene), a polymethyland higher alkylnaphthalene fraction (which is recycled to the reaction Zone) and a bottoms fraction. The naphthalene concentrate is treated in any suitable manner (conveniently by crystallization) to recover the naphthalene in essentially pure form or, if a less pure naphthalene product is acceptable, the na-phthalene distillation column may be operated to produce such `a product directly. If desired, when the naphthalene is recovered by crystallization, the mother liquor may be subjected to further dealkylation either in the reactor of its previous treatment or in a separate dealkylation zone. Also should it not be desired to produce the light gasoline and mother liquor as separate products, the light gasoline and naphthalene concentrate may be produced as a single stream (e.g., by combining gasoline distillation column 10 and naphthalene distillation column 12 of the drawing) and employed directly as'feed to the crystallization operation.

The mono-methylnaphthalene fraction of the reactor eiuent and the process feedstock, as exemplified by the overhead from column 44- of the drawing, typically obtains between about 50 and about 75 weight percent of -Z-methylnaphthalene and between about 10` and about 35 weight percent of l-methylnaphthalene.

While improved processes for the production of naphthalene and 2-rnethylnaphthalene such as those set forth in detail above are preferred embodiments of this invention, this aspect of the invention, in broad concept, contemplates the incorporation of the step of azeotropically distilling non-naphthalenic hydrocarbons from -any mixture containing alkylnaphthalenes and non-naphthalenic hydrocarbons of substantially the same boiling point range, into dealkylating processes in other ways as well. In this connection, alkylnaphthalene heart-cuts from a variety of streams in dealkylation processes lend themselves particularly well to our azeotropic distillation treatment. An example of such a stream is that represented'on the drawing by the mixture of vfeedstock and reactor efliuent in line 3d. Streams of this nature con- -ventionally runl about percent methylnaphthalene and have a 2-methylnaphthalene to l-methylnaphthalene ratio of about 7:2. Another such stream is reactor product of the type exemplified by the material in line 32 on the drawing, which type of material normally runs about 25.5 percent methylnaphthalenes and has a Z-methylnaphthalene to 1-methylnaphthalene ratio of about 30:1. A methylnaphthalene heart-cut from either of these sources normally contains about 90' percent methyl- -naphthalenes and 10 percent of other unidentified materials. In the case of the reactor product the methylnaphthalene heart-cut usually contains about 87 percent Z-methylnaphthalene, whereas the feedstock-reactor effluent heart-cut typically contains about percent 2- methylnaphthalene.

Unless extremely high yields of Z-methylnaphthalene are desired, processing of the reactor product to obtain 2-methylnaphthalene is probably less desirable than processing of the reactor feed (feedstock-reactor efuent mixture) for that purpose. This is true because the reactor product contains much less methylnaphthalene than does the reactor feed (la typical comparison being as indicated above, 25.5 weight percent vs. 65 Weight percent) and more distillation is therefore required to concentrate the methylnaphthalene fraction of the former. Also, all of the methylnaphthalene present in the reactor product has passed through an expensive process step (dealkylation) whereas only a portion of the reactor feed has undergone this reaction step.

Our invention is broad as to the use of suitable azeo trope formers in the Z-methylnaphthalene recovery systern of our dealkylation process, and any material which satisfies the requirements of such an azeotrope former can be so used if desired. In this connection, we have found monoethanolamine to be particularly useful as such an azeotrope former, for reasons given below. Examples of other suitable azeotrope formers are N-methylformamide; butyrolactone; maleic anhydride; lower glycol formates; and the lower alkanolamines other than monoet'hanolamine such as, forexample, 4-amino-2- butanol.v The preferred glycol formates are those lower K substantially all alkylbenzenes, alkyltetralins and alkylindanes boiling within about 25 F. of mono-methylnaphthalenes from the mixtures undergoing treatment, before substantial amounts ofthe mono-methylnaphthalenes themselves distill off. Also, the resulting azeotropic containing from about 20 to about 80 percent by weight VV'of azeotrope former, thus minimizing the distillation load,

and the amount ofazeotroping agent required. Moreover, upon condensing the resulting azeotropic overhead,

twodistinct liquid phases are formed, thus facilitating an obviously benecial recycle of azeotrope former to the distillation column. While," as previously indicated, a small proportion of azeotrope former remains dissolved in the hydrocarbon phase of the condensed overhead, this is easily recovered by an auxiliary azeotropic distillation step, as shown on 14 alkanolamines and butyrolactone are concerned, our invention is c-onsidered broad to the use of those materials as azeotrope formers for the separation of mixtures of alkylnaphthalene and non-naphthalenic hydrocarbons of like boiling point range into fractions of alkylnaphthalenic and non-naphthalenic character.

The .aforesaid mixtures can be from any source and are not limited to dealkylation process, or even the conventional petroleum rening, origins. Non-naphthalenic hydrocarbons which may be present in such mixtures include aliphatic, naphthenic, benzenoid, etc., compounds; and alkylnaphthalenes which may be present include the methylnaphthalenes, and dimethylnaphthalenes, the ethylnaphthalenes, the isopropylnaphthalenes, etc. Attention is called to the fact that copending U.S. patent application Serial No. 19,740 to Backlund, filed April 4, 1960, now abandoned -discloses the use of monoethanolamine as an azeotrope former 4for the separation of non-naphthalenic hydrocarbons from mixtures thereof with naphtha- Iene and that copending application Serial No. 19,526 to McKinnis, filed April 4, 1960, now U.S. Patent No. 3,043,890, discloses the use of butyrolactone for the same purpose. The separation is of course different yfrom the alkylnaphthalene separation of the present invention.

With respect to the above-named azeotrope formers, other than the lower alkanolamines and butyrolactone, broad disclosure of their use of non-naphthalenic hydrocarbon-alkylnaphthalene azeotropy separation can be found in the following copending U.S. patent applications:

Azeotrope Former Inventor (s) Serial Filing Date Patent No. No.

Maleie anhydride McKinuis Flint 135, 624 Sept. 1,1961 3,114, 679 N-methylformamide McKinnis 82, 496 .l an. 13, 1961 3, 10Q, 039 Glycol formates MeKinnis si Webb Sept. 29, 1961 3, 114., 680

the drawing. Thus, no azeotrope former is lost from the system, and no auxiliary extraction systems are required to recoved the azeotroping agent. It is, of course, within the scope of this invention, .and will within the ability of those skilled in the art, to incorporate make-up azeotrope former means in the Z-methylnaphthalene recovery system ofr process, even though such means is normally not requisite thereto. The subject azeotrope formers have good stability under the operating conditions encountered in The preferred proportions of azeotrope former to A,azeotropic distillation feedstock will vary, depending to a large extend upon t-he character of the feedstock. Optimum proportions are ideally those just suicient to form azeotropes with the non-naphthalenic hydrocarbons in the feedstock.

The present invention comprises a unique combination .of process features by means of which two useful products, naphthalene and 2-methylnaphthalene, are atlenic hydrocarbons is novel or not. In this connection, however, we wish to pointl out that the usefulness of certain of the above-disclosed azeotrope formers (the lower alkanolamines, including, as the preferred species,

monoethanolamine, and butyrolactone) for our needs was discovered by us. Accordingly, insofar as the lower Following are examples illustrative of the effectiveness of azeotropy in the Z-methylnaphthalene recovery system of our invention, by contrast with the failure of straight -distillation to achieve the kind of separation there involved. It is cautioned that these examples are included for illustrative purposes only and that they are intended as merely exemplary, and not limitative, of the azeotropic aspects of our invention.

EXAMPLE I This example is included to illustrate the inefficiency of straight distillation as a means of separating alkylnaphthalenes from close boiling non-naphthalenic hydrocarbons in admixture therewith.

The 400w50`0 F. Itraction of reformate product was I separated therefrom by distillation and subjected to a hydrocarbon type analysis. The results of the analysis are set forth below:

Components: Percent by Weight Alkanes 7.6 Mono-naphthenes 2.3 Poly-naphthenes 2.1 Alkylbenzenes 22.1 Tetralin-indanes 24.8 lslaphthalenes1 37.8

1 The distribution of the naphthalenes was found to be as follows (the figures representing Weight percentages of the total sample) CioHs 8.7 CnHro 21.0 CuHiz 6.4 CieHi-t 1.3 CMI-11u 0.3 CisHis 0.1

l 37.3 The 4005001 F. reformate fraction was subjected to batch distillation in a 30-plate Oldershaw column at a reux ratio. A heart-cut fraction of methylnaphthalenes :1 reflux ratio and in the absence of an azeotroping boiling within the range from 456 to 467 F. was obfagent. The fractions obtained from the batch distillation tained. To a -plate Oldershaw column operated at a were analyzed for hydrocarbon types. Results of these 10:1 reflux ratio was charged 200 grams of the aforeanalyses appear in Table 1 below: 5 said nionoethylnapht'nalene heart-cut and 40 grams of Table 1 Boiling point, F. i 406420 l 420-430 430-435 435-440 440-445 445-450 45o-455 45s-460 46o-405 465-470 470-540 540-550 Weight primaria... 18.5 13.3 5. 07 7.14 5. 65 3.38 2. 04 2. 44 4. 22 13 3 14.46 0 40 Aikanes 11.9 9. 8 9.1 8. 7 8.1 7. 2 3. 5 2. 7 0. 7 Moiio-naphthene 4. 1 4. 4 4. 3 4. 5 4. 0 3.3 4. 6 2. 1 1. 4 0` 6 Poiy-naphtheiies.. 1.9 2. 2 2. 7 2. 6 2. 8 3.1 2. 9 2. 3 1.9 0. 7 0. 5 0. 5 Aikyibenzenes 34. 5 34. 1 34. 0 34. 2 32. 2 28. 3 20. i 16. 3 10. 9 4. 4 2. 6 3. 9 Tetraiin-indan5s 23. 9 30. 8 39. 0 38. 9 44. i 48. 5 4i. 7 36. 4 26. 8 9. i 4. 8 1. 5 Naphthaienes 13. 5 19. 2 10.9 11.0 s. 8 9.1 27.1 39.7 58. 3 85.1 80. 6 53. 2 10H5 17.7 18.6 10.2 10.2 i 6.7 5.1 3.3 2.6 1.7 0.7 3.3 3.0 CHH1 0.2 0.4 0.5 0.6 1.8 3.6 23.2 36.5 56.0 83.8 27.3 1.5 C1zHi2. 41.9 11,6 CigHl/l. 6. 5 29. 2 Canam 0 1 0 1 0.6 7.2 CiHis 0.3

In evaluating the Table 1 data, it will be noted that N-methylforrnamide. The overhead (B.P. 204 C.) from no practical separation of the alkylnaphthalenes from the distillation contained two liquid phases; an upper hythe non-naphthalenic compounds present took place iii drocaibon-rich phase and a lower N-methylforrnamidethe higher boiling fractions, particularly those between rich phase. The distillation was continued until all of 450 and 540 F. Thus, while the bulk or" the alkyl- 25 the charged N-inethylformamaide had been collected in naphthalene portion of the feed distilled within that the overhead receiver. The distillate was water washed range, there was much alkybenzene and tetralin-indane and the separated hydrocarbon phase analyzed as was distillation there also to mitigate against good separaa sample of the bottoms. Results are shown in Table tion eflicacy. This was especially true of the boiling 3 bSlOWI point range from 450 to 465 F. 30 Table 3 EXAMPLE II Wt. percent Methylnaphthalene ari- This is an example of the use of monoethanolamine as methylnaphalySiS (Wt. percent) an azeotrope former in the practice of azeotropic distilla- Matemlaualyzed (g') rgresan tion in accordance with the teachings of this invention. 3 2MNaphlMeNaph.

A sample of reactor effluent was subjected to distillaboiling within the range from 456 to 467 F. was obtained. To a ZO-plate Oldershaw column operated a 5:1 40 The Table 3 results clearly evidence the high degree refluX ratio WaS Charged 200 grams 0f the aforesaid of separation attainable between methylnaphthalenes and mOHOmethB/llaphfhalene heart-Cui and 80 ugfams 0f mono' close boiling non-naphthalenic hydrocarbons by the use ethanolamine. rlhe overhead .(BP. 169 C) from the of N-methylformamide as an azeotrope former in the 2- diSlllaUOrl COrlfalIled WO llquld Phases; 21H Upper hY' methylnaphtlialene recovery phase of our invention. drocarbon-rich phase and a lower monoetlianolarnine- 1t Will he apparent that there .are many embodiment rich phase. The distillation was contained until all of variations within the Scope of our inventi0n other the Charged molethanolamll'ie been COllSCed 111 modes of appying the principles 0f the invention may be the overhead receiver. The distillate was water-washed empoyed in place of those Specifically explained, proand tile Separated hYdIOCfbOH Phase Submitted f Of vided any changes with respect to materials or methods analysis. A sample of the bottoms material was likewise fall Within th@ Scope or equivalents of the fouowing submitted for analysis. The results are shown in Table Claims- 2, below: We claim:

T able 2 1. A process for obtaining naphthalene and Z-methylnaphthalene which comprises: Waperceiit Meihyinaphihaiene axi- (1) forming a feed mixture of hydrocarbon compo- Mmeal analyzed (g.) aiysis (wt. percent) nents including methylnaphthalenes and hydrogen;

material (2) passing said feed mixture through a dealkylation 2-1/1 N i N l e aph me aph reaction zone;

(3) withdrawing a reactor effluent from said dealkyliiflag; i i? ation reaction zone and condensign Said @intiem to Bottoms (165) 97.1 9M 3-7 obtain a hydrocarbon liquid eiuent phase and a gaseous efliuent phase;

The Table 2 results clearly evidence the high degree (4) dsiulg Said hydrocarbon liquid efllent Phase of separation attainable between methylnaphthalenes in the absence 0f any Substantial quantity Ovf Said and close boiling non-naphthalenic hydrocarbons by the feed miXtUre I0 reCOVCr a ist fraction comprising use of moiioethanolainine as an azetrope former in the Haphthllelle and 21 Second fraction in which the ratio ZI-methylnaphthalene recovery phase of our invention. Of Z'me'hyhlaijhhalele t0 1methy1lephtha1er1e S EXAMPLE In Zsililstantially higher than that in said feed mixture,

This is an example ofthe use of N-methylformamide (5) azeotropically distilling said second fraction toas an azeotrope former in the practice of azeotropic disgether with an azeotrope former whereby the l10ntillation in 4accordance with the teachings of this invennaphthalenic components together with said azeotion, trope former are vaporized as a minimum boiling A sample of reactor effluent was subiected to distillaazeotropic distillate, thereby leaving as azeotropic ,tion through a 30-plate OlderShaW column at a 10:1 75 bottoms a product enriched in Z-iiiethylnaphthalene.

2. A process for obtaining naphthalene and 2methyl naphthalene which comprises:

(1) forming a feed mixture comprising a hydrocarbon fraction containing alkylnaphthalenes, the hereinafter identified non-naphthalenic hydrocarbon bottoms product, and hydrogen;

(2) passing said feed mixture through a dealkylation reaction zone;

(3) withdrawing a reactor effluent from said dealkylation reaction zone and condensing said eluent to obtain a hydrocarbon liquid effluent phase and a gaseous effluent p-hase;

(4) distilling said hydrocarbon liquid effluent phase in the absence of any substantial quantity of said feed mixture to recover a first fraction which is rich in naphthalene, a second fraction which is rich in mono-methylnaphthalenes and in which the ratio of 2-methylnaphtha1ene to 1methylnaphthalene is substantially higher than that in said feed mixture, and a third fraction comprising alkylnaphthalenes boiling above the mono-methylnaphthalenes;

('5) azeotropically distilling said second fraction together with an azeotrope former whereby the nonnaphthalenic components together with said azeotrope former are vaporized as a minimum boiling azeotropic distillate, thereby leaving as azeotropic bottoms a product enriched in Z-methylnaphthalene;

(6) condensing said azeotropic distillate and allowing the condensate to separate into two liquid phases, one phase comprising said azeotrope former having a minor proportion of non-naphthalenic hydrocarbons and the other phase comprising non-napht-halenic hydrocarbons and a minor proportion of said azeotrope former;

(7) returning the phase comprising the azeotrope former and a minor proportion of non-naphthalenic hydrocarbons to step (8) fractionally distilling the phase comprising nonnaphthalenic hydrocarbons and a minor proportion of the azeotrope former to obtain an overhead azeotropic distillate of azeotrope former and non-naphthalenic hydrocarbons and a bottoms product comprising non-naphthalenic hydrocarbons; and

(9) returning the non-naphthalenic hydrocarbon bottoms product to step (1).

3. The process of claim 2 in which the azeotrope former is monoethanolamine.

4. The process of claim 2 in which t-he azeotrope former is butyrolactone.

5. The process of claim 2 in whi-ch the azeotrope former is maleic anhydride.

6. The process of claim 2 in which the azeotrope former is N-methylformamide,

7. The process of claim 2 in which the azeotrope former is glycol formate.

8. A process for obtaining naphthalene and Z-methylnaphthalene which comprises:

(l) forming a feed mixture comprising a hydrocarbon feedstock containing 1- and Z-methylnaphthalene, the hereinafter identified non-naphthalenic hydrocarbon bottoms product, water vapor and hydrogen, said water vapor being present in an amount corresponding to between about 0.04 and about 0.15 barrel of water per'barrel of hydrocarbon in said feed mixture, and said hydrogen being present in an amount corresponding to between about 4,000 and about 12,000 s.c.f. per barrel of hydrocarbon in said feed mixture;

(2) passing said feed mixture through an elongated reaction zone in contact with a catalyst comprising an alumina carrier having minor amounts of molybdenum oxide and cobalt oxide thereon, said reaction zone being maintained at a temperature between about 950 F. and about 1250 F. and at a pressure between about 600 and about 1200 p.s.i.g.;

(3) introducing into said reaction zone at a plurality of points along the length thereof, a quenching medium selected from the group consisting of hydrogen and water, said quenching medium being employed in an amount sufcient to maintain the ternperature rise along the length of said reaction zone at a value less than about F.;

(4) withdrawing a reactor efliuent from said reaction zone and condensing said effluent to obtain a hydrocarbon liquid effluent phase and a gaseous eluent phase;

(5) distilling said hydrocarbon liquid etiluent phase in the absence of any substantial quantity of said feed mixture to recover a rst fraction which is rich in naphthalene, a second fraction which is rich in monornethylnaphthalenes and in which the ratio of 2- methylnaphthalene to l-methylnaphthalene is substantially higher than that in said feed mixture, and a third fraction comprising alkylnaphthalenes boiling above the mono-methylnaphthalenes;

(6) azeotropically distilling said second fraction in the presence of an azeotrope former whereby substantially al1 of said non-naphthalenic hydrocarbons together with substantially all of said azeotrope former is vaporized as a minimum boiling azeotropic distillate, thereby leaving as azeotropic bottoms a product enriched in 2-methylnaphthalene;

(7) condensing said distillate and allowing the azeotropic condensate t-o separate into two liquid phases, one phase comprising said azeotrope former and a Iminor proportion of non-naphthalenic hydrocarbons and the other phase comprising non-naphthalenic hydrocarbons and a minor proportion of said azeotrope former;

(8) returning the phase comprising the azeotrope former and a minor proportion of non-naphthalenic hydrocarbons to step (6);

(9) fractionally distilling the phase comprising nonnaphthalenic hydrocarbons and a minor proportion of the azeotrope former to obtain an overhead azeotropic distillate of the azeotrope former and nonnaphthalenic hydrocarbons and a bottoms product comprising non-naphthalenic hydrocarbons; and

(10) returning said bottoms product to step (l).

9. A process as defined by claim 8 wherein at least a portion of said third fraction from distillation step (5) is returned to step (l) for employment as an additional component of said feed mixture.

10. A process as dened by claim 8 wherein, in step (5), said distillation is carried out to recover a fourth fraction having a boiling range below that of said' rst fraction.

11. A process as defined by claim 8 wherein said azeotrope former is monoethanolamine.

12. A process as defined by claim 8 wherein said catalyst comprises a carrier comprising alumina containing between about 3 and about 15 percent by weight of silica having distributed thereon between about 5 and about 20 percent by weight of molybdenum oxide, between about 1 and about 5 percent by weight of cobalt oxide, and between about 0.1 and about 3 percent by weight of sodium hydroxide.

References Cited by the Examiner UNITED STATES PATENTS 2,422,673 6/ 1947 Haensel et al. 260-672 3,043,890 7/1962 McKinnis 260-674 3,109,039 10/ 1963 McKinnis 260-674 3,114,679 12/ 1963 McKinnis et al. 260-674 X 3,114,680 12/1963 McKinnis et al. 260-674 X DELBERT E. GANTZ, Primary Examiner.

ALPHONSO D. SULLIVAN, Examiner. 

1. A PROCESS FOR OBTAINING NAPHTHALENE AND 2-METHYLNAPHITHALENE WHICH COMPRISES: (1) FORMING A FEED MIXTURE OF HYDROCARBON COMPONENTS INCLUDING METHYLNAPHTHALENES AND HYDROGEN; (2) PASSING SAID FEED MIXTURE THROUGH A DEALKYLATION REACTION ZONE; (3) WITHDRAWING A REACTOR EFFLUENT FROM SAID DEALKYLATION REACTION ZONE AND CONDENSIGN SAID EFFLUENT TO OBTAIN A HYDROCARBON LIQUID EFFLUENT PHASE AND A GASEOUS EFFLUENT PHASE; (4) DISTILLING SAID HYDROCARBON LIQUID EFFLUENT PHASE IN THE ABSENCE OF ANY SUBSTANTIAL QUANTITY OF SAID FEED MIXTURE TO RECOVER A FIRST FRACTION COMPRISING NAPHTHALENE AND A SECOND FRACTION IN WHICH THE RATIO OF 2-METHYLNAPHTHALENE TO 1-METHYLNAPHTHALENE IS SUBSTANTIALLY HIGHER THAN THAT IN SAID FEED MIXTURE; AND (5) AZEOTROPICALLY DISTILLING SAID SECOND FRACTION TOGETHER WITH AN AZEOTROPE FORMER WHEREBY THE NONNAPHTHALENIC COMPONENTS TOGETHER WITH SAID AZEOTROPE FORMER ARE VAPORIZED AS A MINIMUM BOILING AZEOTROPIC DISTILLATE, THEREBY LEAVING AS AZEOTROPIC BOTTOMS A PRODUCT ENRICHED IN 2-METHYLNAPHTHALENE. 